Two stage hydronitrification and hydrogenation process



March 7, 1967 RF. VAN DRIE SEN 3,308,057

TWO STAGE HYDRONITRIFICATION AND HYDROGENATION PROCESS Filed April 21, 1964 INVENTOR ROGER P. VAN DRIESEN mwrzw ATTORNEY United States Patent 3,308,057 TWO STAGE HYDRONITRIFICATION AND HYDROGENATION PROCESS Roger P. Van Driesen, Hopewell, N.J., assignor to Cities Service Research and Development Company, New

York, N .Y., a corporation of New Jersey Filed Apr. 21, 1964, Ser. No. 361,407 2 Claims. (Cl. 208254) This invention relates to a combination of refining and conversion steps whereby mineral oil fractions boiling in the heavy gas oil range and naphtha range are converted to more desirable products.

Heavy hydrocarbon materials such as residuum material, crude oils, naturally occurring tars and particularly Athabasca bituminous tar require some type of destructive thermal cracking in order to convert the heavy hydrocarbon material to usable distillates. This cracking can be done by various techniques such as coking, hydrovisbreaking, or visbreaking. However, in many cases the thermally cracked distillates obtained from the heavy hydrocarbons are still unsatisfactory for conventional crude oil refining. Thus, the naphthas have a high olefin content, e.-g. a bromine number above 50, whereas the gas oils and particularly heavy gas oils are of low gravity and have high contents of organic nitrogen, e.g. above 0.1% or 0.2% by weight, and high sulfur contents.

It is desirable to furth'erupgrade such distillate-s by hydrotreating. However, due to high nitrogen contents, high sul-fur contents and lowAPI gravities of the heavy gas oil and the high olefin content of the naphtha, difficulties are encountered in the hydrogenation of such distillates. Hydrogenation of these high nitrogen and olefin content distillates cause excessive temperature rise in the reactor and catalyst poisoning. Excess temperature rise is further aggravated in treating the heavy gas oil and naphtha since relatively high temperatures are required in order to obtain nitrogen removal at reasonable hydrogenation severities. In addition to the difficulties caused by exothermic heat of hydrogenation, conventional fixed bed catalyst reactors are poorly suited for treating the heavy gas oil of low API gravity and high nitrogen content since the heavy gas oil causes fouling and plugging of the catalyst bed. Also, the heat of hydrogenation for the heavy gas oil and of the naphtha is difiicult to remove in a fixed catalyst bed reactor and further it is difiicult to maintain substantially isothermal conditions throughout a fixed catalyst bed when substantial quantities of heat are liberated in hydrogenation.

It is an object of this invention to provide a process for upgrading both naphtha having a high olefin content and heavy gas oil having a low API gravity and high nitrogen content.

It is another object of this invention to provide a multistage process for hydrogenating heavy gas oil wherein the exotherm from the first stage is removed by quenching the efiluent with addition of naphtha and subsequently hydrogenating both the upgraded heavy gas oil efiluent and the naphtha.

It is still another object of this invention to provide a process for treating heavy gas oil of low API gravity and high nitrogen content and naphtha of high olefin content in a process which obtain-s maximum usage of the catalyst mass, exothermic heat of hydrogenation and wherein the catalyst mass is maintained under substantially isothermal conditions.

It is a further object of this invention to provide a process for hydrotreating a highly unsaturated thermal naphtha in contact with a substantially isothermal ebullated catalyst bed while the naphtha is in the vapor phase "ice Briefly, the process of this invention comprises operating a series of two or more hydrogenation reactors wherein the heavy gas oil is partially hydrogenated in the first reactor. Between reactors the naphtha is injected into the heated and partly hydrogenated gas oil. The naphtha is injected into this effluent for the most part in the liquid phase at a temperature below that of the first stage efiluent. This causes considerable cooling of the efliuent stream both by the sensible heat required to bring the naphtha up to the temperature of the feed stream and the latent heat required to vaporize the naphtha. The cooled mixture of the partially hydrogenated gas oil and naphtha is then fed into another reactor wherein the mixture is hydrogenated. In the mixture of naphtha and cooled heavy gas oil efiiuent and in the subsequent hydrogenation of the mixture, the gas oil effluent is for the most part in the liquid phase whereas the naphtha is mostly in the vapor phase. The process affords means within highly effective temperature ranges for hydrogenating heavy gas oil of low API (American Petroleum Institute) gravity and high nitrogen contents, cooling the hydrogenated efiiuent between reaction stages, preheating the naphtha and hydrogenating naphtha of high olefin content in the vapor phase together with the partially hydrogenated gas oil in the liquid phase. Also, the process provides liquid hydrocarbons in each stage of hydrogenation which permits the use of an ebullated catalyst bed. Thus, having the problem of hydrotreating a highly unsaturated thermal naphtha an ebullated bed catalyst mass maintained under substantially isothermal conditions can be employed, even though the naphtha is in the vapor phase, hydrotreating the naphtha along with heavy gas oil. Hydrotreating both the naphtha and heavy gas oil by the process of this invention does not require any more reactor volume than that required for hydrotreating the heavy gas oil without the naphtha. In fact the heavy gas oil hydrotreating operation is improved, considering a fixed reactor size and all other factors constant, by the blending of the naphtha. The two stage hydrotreating of the heavy gas oil greatly aids in nitrogen and sulfur removal and since there is a cooling requirement between stages, the addition of naphtha is advantageously made between stages to provide the cooling. The term substantially isothermal conditions as employed herein in describing the temperature of the catalyst mass is intended to mean temperature differential of less than about 10 F. and preferably less than about 5 F. throughout the catalyst mass or bed.

For the most part, the nitrogenous components of the hydrocarbon oil feeds of this process are reduced to ammoma.

The efiduent from this multistage process can then be taken to fractionation where the naphtha is flashed off and processed by further hydrogenation.

The ratio naphtha injected into the partially hydrogenated heavy gas oil efiluent can vary over a wide range such as that from about 0.1 volume to about 2 volumes of naphtha per volume of treated gas oil effluent and preferably from about 0.2 to 1.5 volumes of naphtha per volume of the treated gas oil efiluent.

Preferably the heavy gas oil employed in this invention has an API gravity of from about 5 to about 20 API; a nitrogen content of from about 0.1 or 0.15% to about 0.7% by weight; a sulfur content from about 2% to about 7% by weight; and particularly an API gravity from about 6 to about 18 API; a sulfur content of from about 5% to about 6%; and a nitrogen content of w from about 0.2% to about 0.7%.

The naphtha employed in this process has a substantial 'unsat'uration with a bromine number of at least 50 such that of from about 50 to about 140 and particularly a bromine number in excess of about 80. The naphtha will also generally have a nitrogen content of about to 400 parts per million (p.p.m.), by weight, and a sulfur content of about 1% to about 3% by weight.

In each stage of the hydrogenation the nitrogen in the efiluents boiling in the heavy gas oil range can be reduced by about 10% to about 80% of the amount present in the heavy gas oil feed or the feed of treated gas oil efiluent and more usually from about 30% to about 50% of the amount present in the gas oil feed to each stage. Total nitrogen reduction in the entire process of effluent from the last stage boiling in the havy gas oil range can be from about 30% to about 95% of the amount of nitrogen present in the heavy gas oil feed. Sulfur content of the heavy gas oil effluent in the process is decreased by about 50% to about 95% of its original value and more usually from about 70% to about 80%. The API gravity of material boiling in the heavy gas oil range is raised by from about 2 to 15 in the process and more usually from 7 to about 12 with about a 2 to 10 API gravity change, and more usually a 5 to 7 change occurring in the first stage.

The effluent recovered from the process in the naphtha range has a bromine number of from about 5 to 40 and particularly about 10, generally representing a reduction of 70% to 90% of the unsaturation based on bromine number. Also, the nitrogen content of the naphtha can be reduced generally to about 50 ppm, although a range of from 10 to about 100 p.p.m. is satisfactory. The reduction of nitrogen in the naphtha being generally about 50%-90% by weight.

The heavy gas oil employed in this invention can be one having a boiling point of from about 650 F. to about 1100 F. Whereas the naphtha has a boiling point from about 80 F. to about 380 F.

In a preferred embodiment of the invention the hydrogenation catalyst mass or bed is expanded by the upward flow of fluids in the reactor, e.g. an ebullated bed as described in U.S. Patent 2,987,465 to E. Iohanson which issued on June 6, 1961. The use of an expanded catalyst mass facilitates the maintenance of substantially isothermal operations in this invention. Since a liquid is required in the ebullated catalyst bed, the naphtha, which is in the vapor phase at the temperature employed, is passed through the catalyst bed together with a heavy gas oil effluent in the liquid phase. In the use of an ebullated catalyst mass or bed, the gross volume of the mass expands without, however, a substantial quantity of the catalyst particles being carried away by the upfiowing fluids, and, therefore, a fairly well defined upper level of randomly moving particles establishes itself in the upflowing fluids of the reactor. The fluid flow rate in each reactor when employing an ebullated bed can be easily controlled with a conventional space velocity and recirculation of fluids to prevent catalyst particles from being carried away.

In place of an ebullated bed of catalyst, a catalyst in the form of more finely divided particles can be advantageously employed in the form of a slurry in the hydrocarbon oil. A slurry catalyst is of sufficiently small particle size so that it can be carried along with the hydrocarbon oil instead of forming a fairly well defined upper level in the reactor. A combination of both ebullated bed and slurry catalyst systems can also be employed. The ebullated bed and slurry catalyst systems are referred to herein as an expanded catalyst mass.

The hydrogenation catalysts can be in the form of beads, chips, pellets, lumps, or the like. The particle size of the catalyst can vary over a wide range. Illustratively, the average particle size of ebullated bed catalyst can be from about of an inch to about A of an inch or more. When a slurry of catalyst in the oil is employed, the catalyst can have a particle size of less than about 300 .microns with that of about 50-150 microns preferred.

The term particle size as used herein refers to an average dimension of a catalyst particle.

The term hydrogenation as employed herein is intended to include any reaction where hydrogenation takes place irrespective of additional reactions such as cracking. Thus the hydrogenation catalyst employed can be one which simply promotes hydrogenation or one which promotes both hydrogenation cracking, e.g. a dual function catalyst.

Any hydrogenation or hydrocracking catalyst may be used in the reactors. Satisfactory results can be obtained with platinum, palladium, molybdenum, iron, nickel cobalt, tungsten and the like. In addition, any such suitable metal or its oxide or sulfide may be used in combination with inert surfactive carrier or conventional acidic cracking material such as silica-alumina, silica-magnesia, silicaalurnina-zircom'a, acid treated clays and the like. For maximum utilization of a given reactor when employing the ebullated bed technique, catalyst generally is used in quantities sufiicient to provide an average concentration of at least 15 and preferably at least 25 pounds per cubic foot of contact zone when in the ebullated state.

Each of the hydrogenation reaction vessels may be maintained under suitable pressure such as between about 500 and 4,000 p.s.i.g. partial pressure of hydrogen with between about 1,000 and 2,000 p.s.i.g. partial pressure of hydrogen being preferred.

Temperature in each of the reaction zones can be maintained within desired limits within a broad range such as that of from about 650 F. to about 950 F., depending upon the particular hydrogenation reaction desired and other operating conditions with temperatures between about 700 F. to about 850 F. being preferred. Substantially isothermal conditions are preferably maintained in each of the catalyst beds by recirculation of the liquid hydrocarbon oil in each reactor. In the reactor containing both heavy gas oil eifluent and vapor phase naphtha, isothermal conditions are maintained by recirculation of heavy gas oil efiluent in the liquid phase. The quantity of liquid recycle may vary widely such as between about 1 and about 60 times the volume of liquid feed introduced in a reactor. Recycle rates of from between about 5 to about 25 volumes of recycle liquid per volume of fresh feed are generally preferred. A vertical open ended draft tube can be used to affect recirculation or recycle of liquids in each reactor. The temperature of the feed to each reactor is preferably from about 20 F. to about F. lower than that of the temperature in the reactor.

The naphtha is added to the heavy gas oil efiluent at a temperature substantially lower than that of the effiuent. Also, a major or appreciable portion of the naphtha is in the liquid phase prior to contact with the gas oil efiluent. Illustratively the naphtha can be at a temperature of from about 50 F. to 200 F. below that of the partially hydrogenated gas oil effluent.

Hydrogen feed into the reaction zones can be substantially pure hydrogen or part of a hydrogen-containing gas, such as recycle gas containing hydrogen. Hydrogen is introduced into each of the reactors in quantities suitable for the hydrogenation desired in each of the reactors and may for instance be introduced at rates between about 1,000 to about 30,000 standard cubic feet of hydrogen per barrel of hydrocarbon feed with hydrogen rates between about 2,500 and 10,000 s.c.f. per barrel of feed being preferred. The hydrogen can be introduced separately into each reactor, such as by being introduced together with the heavy gas oil in one reactor and together with the naphtha into the gas oil effluent stream prior to hydrogenation of the gas oil-naphtha mixture in a subsequent reactor, or sufficient hydrogen can be introduced with the heavy gas oil feed for the entire series of reactors.

The space velocity in each reactor can vary widely such as between about 0.1 to about 5 volumes of feed per hour per volume of reactor capacity and preferably from about 1.0 volume of feed per hour per volume of reactor capacity.

For a better understanding of the invention reference is made to the accompanying drawing wherein identical numerals refer to identical parts and wherein reactor 12 has an inlet line 10 for admitting heavy gas oil and hydrogen which pass upwardly in reactor 12 through opening 16 of grid 14. Caps 18 over grid openings 16 further limit the passage of the upflowing fluids and prevent catalyst particles from clogging the openings 16. The heavy gas oil feed and hydrogen pass upwardly through the catalyst bed 28, expand the volume of the catalyst bed and the gas oil is hydrogenated in contact with the catalyst. The fluids in the reactor continue their upward movement and product is recovered at the top of reactor 12 through open valve 30 and line 26. A major portion, by weight, of hydrocarbon effluent from reactor 12 is in the liquid phase. A portion of the liquid in reactor 12 is recycled by passing downwardly through funnel top 24 and out of opening 22 of draft tube 20. This maintains substantially isothermal conditions throughout the entire catalyst bed area. The temperature differential from the bottom of the catalyst bed to the top of the catalyst bed can thus be controlled within less than about 10 F. and preferably within less than about F. Naphtha, a major portion of which is in the liquid phase, optionally with additional hydrogen is added to the gas oil effluent in line 26 through line 32 and valve 34. The mixture of naphtha and gas oil efiluent in line 26 is fed into reactor 40 through opening 38. The incoming liquids together with gases pass upwardly through opening 44 in grid 42 and catalyst mass 58 is again expanded and the feed from line 38 is hydrogenated in contact with the catalyst mass 58. Caps 46 prevent clogging of grid openings 44. Product is removed from the top of reactor 40 by line 54 and valve 56. A portion of the liquid in reactor 40 is recycled by passing downwardly through funnel top 52 and bottom opening 50 of draft tube 48. Again, the recycling maintains substantially isothermal conditions throughout the entire catalyst bed area. The following example, taken with the drawing, is illustrative of the invention.

EXAMPLE 1 Heavy gas oil obtained from coking Athabasca bituminous tar is fed into reactor 12 at a temperature of 65 0 F. through line together with 4,000 standard cubic feet of hydrogen per barrel of heavy gas oil. A hydrogen partial pressure of 1200 p.s.i.g. is maintained in reactor 12. The gas oil is fed at a rate of 46,630 barrels per day and has a boiling point of from about 650 F. to about 925 R, an API gravity of 15.4, a 4.5% sulfur content by weight and a nitrogen content of 3,000 ppm. by weight. The reactor 12 and catalyst mass or bed 28- is maintained at about 750 F., a space velocity of the volume .of heavy gas oil fed per hour per volume of reactor capacity is maintained at about 2.0. The gas oil is hydrogenated in contact with the catalyst which has a particle size of 4 diameter by A" long and is a composite of cobalt-molybdate on alumina. The catalyst bedis expanded by about 50% of its settled volume due to the passage of fluids upwardly through the reactor. A recycle rate of about ten times that of the volume of liquid feed is maintained by draft tube 20. Partially hydrogenated gas oil product is recovered from reactor 12 through valve 30 and line 26. The effluent from reactor 12 is at a temperature of 750 F. Naphtha at a rate of 21,114 barrels per day and a temperature of 470 F. from line 32 and valve 34 is added to the efiluent in line 26. A major portion of the naphtha in line 32 is in the liquid phase. The naphtha cools the efiluent from reactor 12 and the mixture of reactor 12 efiluent and naphtha has a temperature of about 660 F. This mixture then enters reactor 40 through opening 38. The reactor 40 is maintained at a temperature of about 750 F. The feed into the reactor 40 is hydrogenated in much the same manner as in reactor 12, namely by passing upwardly through openings 44 in grid 42 between the edges of the opening and caps 46 expanding the catalyst mass 58 by about 50% of its settled volume wherein the hydrocarbon fluids are hydrogenated. In reactor 40 the material boiling in the heavy gas oil range is in the liquid phase whereas that boil- 6' ing in the naphtha range is in the vapor phase. Recircu lation is provided in reactor 40 by draft tube 48 inthe same manner as in reactor 12, and the temperature throughout the catalyst mass is maintained at about 750 F. with less than 5 F. difference in temperature throughout the mass. The products are recovered from line 54 and valve 56 from reactor 40.

EXAMPLE 2 Naphtha and heavy gas oil from a thermal visbreaking operation can be hydrotreated in the apparatus of the accompanying drawing in much the same manner described in Example 1. Heavy gas oil having a boiling point of 650 F. to 925 F. is fed into reactor 12 at a rate of 40,452 barrels per day. Naphtha at a rate of 17,898 barrels per day together with C hydrocarbons from the visbreaker gas at a rate of 1,910 barrels per day are fed into line 26 from line 32. The following operating conditions are used in each of the reactors 12 and 40: total pressure of 1500 pounds per square inch gauge (p.s.i.g.) with a hydrogen partial pressure of 1200 p.s.-i.g.; temperature of 750 F.; a space velocity of 1.0 based on volume of heavy gas oil per hour per reactor volume; and a hydrogen consumption of 960 standard cubic feet per barrel of feed. The catalyst mass in each reactor is maintained under substantially isothermal conditions by recirculation of liquid through the draft tubes in each reactor. The naphtha and C hydrocarbons introduced into line 26 from line 32 is at a temperature of 460 F. The catalyst employed is the same as that of Example 1 whereas a recycle rate of about 12 times that of the volume of liquid feed is employed in each of draft tubes 20 and 48. The yields, based on percent of feed for the process are given in Table 1. The quality characteristics of the feed and of the product are given in Table 2. The symbols C C C and C refer to hydrocarbons having 1, 2, 3 and 4 carbon atoms respectively.

Table 1.Yields based on percent of feed H S, weight percent 3.70 NH weight percent 0.15 C weight percent 0.50 C weight percent 0.40 C weight percent 0.40 C weight percent 3.60 Naphtha, volume percent 36.60 Hydrocarbons boiling in the 380650 F. range,

volume percent 11.00 Hydrocarbons boiling above 65 0 F., volume percent 51.73

Table 2.Quality of liquid Feed Product Naphtha:

(a) Gravity, API 51.1 56 (b) Sulfur, weight percent 2. 28 0.25 (c) Nitrogen, p.p.m 50 (d) Bromine Number 8 Hydrocarbons Boiling in the 380 F. to 650 F.

Range:

(9.) Gravity, API 26.0 (b) Sulfur, weight percent 0. 3 (c) Nitrogen, p.p.m 500 (d) Diesel Index 29. 0 Hydrocarbons Boiling above 650 F.:

(a) Gravity, API 13. 8 20. 0 (b) Sulfur, weight percent 4. 72 0.85 (c) Nitrogen, p.p.m 2,900 1,600

What is claimed is: 1. Process for hydrogenating hydrocarbons which comprises:

(a) feeding an admixture of hydrogen and a heavy gas oil feed boiling in the range of from about 650 F. to about 1100 F. and containing at least about 0.1% by weight of organic nitrogen into a first hydrogenation zone upwardly through a mass of hydrogenation catalyst contained therein at a rate suflicient to maintain the catalyst mass in a random motion expanded state; but below that at which substantial quantities of catalyst are raised to an upper portion of said hydrogenation zone, said catalyst mass being maintained under substantially isothermal conditions;

(b) withdrawing a first heated efiluent of said hydrogenated oil from an upper portion of said first hydrogenation zone wherein said hydrogenated oil eflluent contains a weight proportion of organic nitrogen that is between about 10% and about 80% of that contained in said heavy gas oil feed;

(c) mixing said heated efiluent with a naphtha hydro- (d) feeding hydrogen and said admixture of said first efiluent and said naphtha into a second hydrogenation zone upwardly through a catalyst mass contained therein at a rate sufiicient to maintain the catalyst mass in a random motion expanded state, but below that at which substantial quantities of catalyst are raised to an upper portion of said hydrogenation zone, said catalyst mass being maintained under substantially isothermal conditions, and said first effiuent being maintained substantially in the liquid phase while said naphtha is maintained substantially in the vapor phase in said second hydrogenation zone; and

(e) withdrawing an effluent wherein the bromine number of the naphtha range hydrocarbons has been reduced by about 70% to about 90% compared to the naphtha feed and wherein the organic nitrogen content of the first efiluent has been reduced by about 10% toabout 80%.

,2. Process as in claim 1 wherein hydrocarbon liquid is recycled within each hydrogenation zone to maintain the catalyst mass in each hydrogenation zone under substantially isothermal conditions.

References Cited by the Examiner UNITED STATES PATENTS 3/1959 Hennig 2082l0 2,901,417 8/1959 Cook et al 20*8210 2,987,465 6/1961 Johanson 2O8-10 DELBERT E. GANTZ, Primary Examiner.

SAMUEL P. JONES, Examiner. 

1. PROCESS FOR HYDROGENATING HYDROXCARBONS WHICH COMPRISES: (A) FEEDING AN ADMIXTURE OF HYDROGEN AND A HEAVY GAS OIL FEED BOILING IN THE RANGE OF FROM ABOUT 650*F. TO ABOUT 1100*F. AND CONTAINING AT LEAST ABOUT 0.1% BY WEIGHT OF ORGANIC NITROGEN INTO A FIRST HYDROGENATION ZONE UPWARDLY THROUGH A MASS OF HYDROGENATION CATALYST CONTAINED THEREIN AT A RATE SUFFICIENT TO MAINTAIN THE CATALYST MASS IN A RANDOM MORTION EXPANDED STATE; BUT BELOW THAT AT WHICH SUBSTANTIAL QUANTITIES OF CATALYST ARE RAISED TO AN UPPER PORTION OF SAID HYDROGENATION ZONE, SAID CATALYST MASS BEING MAINTAINED UNDER SUBSTANTIALLY ISOTHERMAL CONDITIONS; (B) WITHDRAWING A FIRST HEATED EFFLUENT OF SAID HYDROGENATED OIL FROM AN UPPER PORTION OF SAID FIRST HYDROGENATION ZONE WHEREIN SAID HYDROGENATED OIL EFFLUENT CONTAINS A WEIGHT PROPORTION OF ORGANIC NITROGEN THAT IS BETWEEN ABOUT 10% AND ABOUT 80% OF THAT CONTAINED IN SAID HEAVY GAS OIL FEED; (C) MIXING SAID HEATED EFFLUENT WITH ANAPHTHA HYDROCARBON BOILING IN THE RANGE OF FROM ABOUT 80%F. TO ABOUT 380*F. AND HAVING A BROMINE NUMBER OF AT LEAST ABOUT 50, SAID NAPHTHA BEING AT A LOWER TEMPERATURE THAN SAID HEATED EFFLUENT, AND AT LEAST A PORTION OF SAID NAPHTHA ADMIXED WITH THE EFFLUENT BEING IN THE LIQUID PHASE; 